Process for the continuous hydrogenation of carbon-carbon double bonds in an unsaturated polymer to produce a hydrogenated polymer

ABSTRACT

Proposed is a process for the continuous hydrogenation of carbon-carbon double bonds in an unsaturated polymer to produce a hydrogenated polymer, said unsaturated polymer being based on a conjugated diolefin and at least one other copolymerizable monomer, in the presence of a homogeneous or heterogeneous catalyst, wherein said unsaturated polymer, hydrogen and said homogeneous or heterogeneous catalyst is passed through a multistage agitated reactor, comprising a cylindrical, elongated shell having closed ends and separated by baffles into a multiplicity of discrete chambers with access from one chamber to another through concentric circular openings, axially centered with said baffles and a continuous rotatable shaft extending concentric with said baffles within said shell with at least one impeller attached thereto positioned in each chamber, said continuous rotatable shaft and said circular openings providing annular openings and said hydrogenated polymer being withdrawn at the opposite end of the multistage agitated reactor at which the feed is introduced.

FIELD OF THE INVENTION

The invention is directed to a process for the continuous hydrogenationof carbon-carbon double bonds in an unsaturated polymer with hydrogen toproduce a hydrogenated polymer, said unsaturated polymer being based ona conjugated diolefin and at least one other copolymerizable monomer, inthe presence of a solvent and a catalyst, preferably a homogeneouscatalyst.

BACKGROUND OF THE INVENTION

The residual unsaturated bonds in polymer backbones are susceptible tobreakdown when exposed to heat, light or ozone. By hydrogenating theseunsaturated bonds, the performance of the material with respect toresistance to heat and ozone can be significantly improved, as well asits durability over long term exposure to aggressive environments.

The process of hydrogenation of unsaturated polymers could be operatedin a batch or a continuous manner. It has been known that the batchprocesses are sometimes expensive, lengthy (a process cycle consists ofpolymer and catalyst preparation, the setting of reaction conditions,the actual reaction time, followed by the cooling and depressurizationof the reactors and the removal of the product) and labour-intensive,and are only suitable for the production of small quantities. Productconsistency is difficult to attain. When used for large productionvolumes, the batch process would require a very large reaction apparatusand very long cycle time. In contrast, when a continuous process isadopted, a large quantity of product can be obtained with consistentqualities and a reactor of a smaller size can be utilized. In addition,the integration of mass balance and heat balance can be realized througha continuous process.

Historically, a continuous process for the hydrogenation of unsaturatedpolymers generally was carried out in fixed bed reactors, wherein thereactors were packed with various types of heterogeneous catalysts.

U.S. Pat. No. 6,395,841 discloses a continuous process for thehydrogenation of unsaturated aromatic polymers in a fixed bed reactorusing a group VIII metal as catalyst. However, a relatively high gasflow rate/polymer solution flow rate ratio (at least 150, vol/vol), wasused. The high flow ratio ensured enough hydrogen transfer from hydrogengas phase to liquid phase. However, the excess of hydrogen gas has to berecycled by a compressor, leading to an increase in operation cost.Furthermore, because of the presence of the packing in the fixed bedreactor, the ratio of the reactor space that can be used for reaction islow.

U.S. Pat. No. 5,378,767 discloses a method of hydrogenating unsaturatedpolymers with low molecular weight which may contain functional groupssuch as hydroxyl in a fixed bed wherein the reactor is packed withplatinum, palladium or a mixture of the two catalysts supported on analpha alumina support. This process may not be suited to handle highmolecular weight polymers.

U.S. Pat. No. 6,080,372 discloses a continuous stirred tank reactorpaired with a bubble column reactor to enhance conversion in acontinuous hydrogenation process. Raney nickel catalyst was used tohydrogenate glucose to sorbitol and a reasonably high hydrogenationdegree (over 90%) can be achieved. However, the use of a bubble columnrequires an excess amount of hydrogen to mix gas and liquid and providesufficient gas-liquid contact.

A multistage agitated contactor (MAC) has been recognized to have manyadvantages over traditional single stage agitated contactors and bubblecolumn reactors. Only a few applications of MACs as gas liquidcontactors have been reported. However, most of the applications arefocused on air-water systems at ambient conditions. Regarding the use ofa MAC for industrial applications in a continuous manner, very few caseshave been reported. U.S. Pat. No. 4,275,012 by Kokubo et al. disclosedits use for the continuous process of refining oils and fats. U.S. Pat.No. 4,370,470 by Vidaurri et al. disclosed its use for the continuousproduction of arylene sulphide polymer in a MAC. However, only a liquidphase reaction was involved and a gas phase reactant was not involved inthe above mentioned processes.

In the prior art, it can be seen that only fixed bed reactors have beenemployed for the continuous process for hydrogenation of unsaturatedpolymers, particularly, using heterogeneous catalysts. However, somevaluable polymers are obtained by using more efficient and highlyselective homogeneous catalysts. For example, Rempel, G. L., 2000,Catalytic Hydrogenation of Nitrile Butadiene Rubber, Polymer Preprints,41(2), 1507 reported several effective catalysts for selectivehydrogenation of nitrile butadiene rubber, including rhodium, rutheniumand osmium catalysts. However, these homogeneous catalyst systems areconducted in batch processes.

When fixed bed reactors are applied to perform the continuous process,it is not economical due to the low applicable reactor volume ratio andhigh pressure drop. If a single continuous stirred tank reactor (CSTR)is considered, an extremely large reactor is needed for a long reactiontime since a high hydrogenation degree (95%) is usually required for theproduction of the final polymer. A continuous process for polymerhydrogenation has some special requirements such as instantaneous mixingof the catalyst at the inlet, exothermic peak mitigation and backflowprevention between the two stages; therefore, the existing MACsmentioned above are not applicable for polymer hydrogenation.

SUMMARY OF THE INVENTION

Accordingly, it is the objective of the present invention to provide aprocess suitable for continuous hydrogenation of unsaturated polymers,wherein the hydrogenation rate is high, steady stable performance isachieved and a flexible operation is possible.

This objective is achieved by a process for the continuous hydrogenationof carbon-carbon double bonds in an unsaturated polymer to produce ahydrogenated polymer, said unsaturated polymer being based on aconjugated diolefin and at least one other copolymerizable monomer, inthe presence of a catalyst, preferably a homogeneous catalyst, whereinsaid unsaturated polymer, hydrogen and said catalyst are passed througha multistage agitated reactor.

BRIEF DESCRIPTION OF THE DRAWING

FIG. 1 shows an example of the multistage agitated reactor.

DETAILED DESCRIPTION OF THE INVENTION

In one embodiment of the invention the process is performed using amultistage agitated reactor comprising a cylindrical, elongated shellhaving closed ends and separated by baffles into a multiplicity ofdiscrete chambers with access from one chamber to another throughconcentric circular openings, axially centered within said baffles andone or two continuous rotatable shafts extending concentric with saidbaffles within said shell with at least one impeller attached theretopositioned in each chamber, and said hydrogenated polymer is withdrawnat the opposite end of the multistage agitated reactor at which the feedis introduced.

In a preferred embodiment of the invention the process is performed withsaid multistage agitated reactor in which the aforementioned continuousrotatable shaft and said circular openings provide annular openings

The present invention provides a continuous reactor for efficientlyhydrogenating unsaturated polymers. The reactor has a high length todiameter ratio and can sustain high pressure having a thinner wallcompared to a CSTR. Mechanical agitation in each stage provides superiormass transfer as well as heat transfer since the hydrogenation reactionsare usually exothermic. This reactor can be operated at a low hydrogenflow rate/polymer solution flow rate ratio. Another advantage associatedwith the reactor in the present invention is that each stage can beoperated at different conditions such as different temperatures withdifferent impellers and even different agitation speeds, depending onthe requirements for mixing and mass transfer performance.

The polymers which are hydrogenated according to the present continuousprocess are polymers which contain carbon-carbon double bonds and whichare based on a conjugated diolefin and at least one othercopolymerizable monomer.

The conjugated diolefin is preferably one or more substances selectedfrom butadiene, isoprene, piperylene and 2,3-dimethylbutadiene,preferably butadiene and/or isoprene, and most preferably butadiene.

At least one other copolymerizable monomer is preferably one or moresubstances selected from acrylonitrile, propyl acrylate, butyl acrylate,propyl methacrylate, methacrylonitrile, butyl methacrylate and styrene,and preferably acrylonitrile and styrene.

Further examples of suitable monomers are esters of ethylenicallyunsaturated mono- or dicarboxylic acids such as acrylic acid,methacrylic acid, maleic acid, fumaric acid, itaconic acid and mesaconicacid with generally C₁-C₁₂ alkanols, such as methanol, ethanol,n-propanol, isopropanol, 1-butanol, 2-butanol, isobutanol,tert.-butanol, n-hexanol, 2-ethylhexanol, or C₅-C₁₀ cycloalkanols, suchas cyclopentanol or cyclohexanol, and of these preferably the esters ofacrylic and/or methacrylic acid, examples being methyl methacrylate,n-butyl methacrylate, ter-butyl methacrylate, n-butyl acrylate,2-ethylhexyl acrylate and tert butyl acrylate.

The hydrogenation of the polymer is preferably undertaken in solution.Especially suitable solvents for the polymer and the hydrogenationprocess include benzene, toluene, xylene, monochlorobenzene andtetrahydrofuran, with monochlorobenzene and tetrahyrodofuran beingpreferred and monochlorobenzene being most preferred. The concentrationof the unsaturated polymer in the solvent may be from about 1 to about40 wt.-%, preferably from about 2 to about 20 wt.-%.

The hydrogenation is undertaken in the presence of a homogeneous orheterogeneous catalyst, which preferably is an organo-metallic catalyst,most preferred a rhodium, ruthenium, titanium, osmium, palladium,platinum, cobalt, nickel or iridium either as metal or preferably in theform of metal compounds (cf., for example, U.S. Pat. No. 3,700,637,DE-A-25 39 132, EP-A-0 134 023, DE-A-35 41 689, DE-A-35 40 918, EP-A-0298 386, DE-A-35 29 252, DE-A-34 33 392, U.S. Pat. No. 4,464,515 andU.S. Pat. No. 4,503,196).

In one preferred embodiment the homogeneous catalyst represents anorgano-metallic catalyst, preferably a rhodium, ruthenium, osmium, oriridium metal complex catalyst.

Preferred metals for the heterogeneous catalyst are one or more metalsselected from platinum, palladium, nickel, copper, rhodium andruthenium. The heterogeneous catalyst can be preferably supported oncarbon, silica, calcium carbonate or barium sulphate.

Preferably, the catalyst is a homogeneous catalyst.

Specially suited are osmium catalysts having the formulaOs QX(CO)(L)(PR₃)₂in which Q may be one of hydrogen and a phenylvinyl group, X may be oneof halogen, tetrahydroborate and alkyl- or aryl-carboxylate, L may beone of an oxygen molecule, benzonitrile or no ligand, and R may be oneof cyclohexyl, isopropyl, secondary butyl and tertiary butyl saidtertiary butyl being present only when one R is methyl, with the provisothat when Q is phenylvinyl X is halogen and L is no ligand and when X isalkyl- or aryl-carboxylate Q is hydrogen and L is no ligand, saidhalogen being selected from chlorine and bromine. Preferably, Q ishydrogen, X is selected from chlorine, tetrahydroborate and acetate, Lis an oxygen molecule or no ligand and R is cyclohexyl or isopropyl.Additional alkyl- or aryl-carboxylates include chloroacetate andbenzoate.

Examples of suitable osmium catalysts includeOsHCl(CO)[P(cyclohexyl)₃]₂, OsHCl(CO)[P(isopropyl)₃]₂,OsHCl(O₂)(CO)[P(cyclohexyl)₃]₂, OsHCl(O₂) (CO)[P(isopropyl)₃]₂,Os(CH═CH—C₆H₅) Cl(CO)[P(cyclohexyl)₃]₂, Os(CH═CH—C₆H₅)Cl(CO)[P(isopropyl)₃]₂, OsH(BH₄)(CO)[P(cyclohexyl)₃]₂, OsH(BH₄)(CO)[P(isopropyl)₃]₂, OsH(CH₃COO)(CO)[P(cyclohexyl)₃]₂, OsH(CH₃COO)(CO)[P(isopropyl)₃]₂, OsHCl(CO)(C₆H₅CN) [P(cyclohexyl)₃]₂, and OsHCl(CO)(C₆H₅CN) [P(isopropyl)₃]₂. Preferred catalysts are OsHCl (CO)[P(cyclohexyl)₃]₂, OsHCl(CO) [P(isopropyl)₃]₂, OsHCl(O₂)(CO)[P(cyclohexyl)₃]₂ and OsHCl (O₂)(CO) P(isopropyl)₃]₂.

The quantity of the osmium catalyst required for the hydrogenationprocess is preferably from about 0.01 to about 1.0 wt.-% based on theunsaturated polymer and most preferably from about 0.02 to about 0.2wt.-% based on the unsaturated polymer.

The selective hydrogenation can be achieved, for example, in thepresence of a rhodium- or ruthenium-containing catalyst. It is possibleto use, for example, a catalyst of the general formula(R¹ _(m)B)₁ M X_(n)where M is ruthenium or rhodium, the groups R¹ are identical ordifferent and are each a C₁-C₈-alkyl group, a C₄-C₅-cycloalkyl group, aC₆-C₁₅-aryl group or a C₇-C₁₅-aralkyl group. B is phosphorus, arsenic,sulphur or a sulphoxide group S═O, X is hydrogen or an anion, preferablyhalogen and particularly preferably chlorine or bromine, 1 is 2, 3 or 4,m is 2 or 3 and n is 1, 2 or 3, preferably 1 or 3. Preferred catalystsare tris(triphenylphosphine)rhodium(I) chloride,tris(triphenylphosphine)rhodium(III) trichloride and tris(dimethylsulphoxide)rhodium(III) trichloride and alsotetrakis(triphenylphosphine)rhodium hydride of the formula (C₆H₅)₃P)₄RhHand the corresponding compounds in which the triphenylphosphine has beencompletely or partly replaced by tricyclohexylphosphine. The catalystcan be utilized in small amounts. An amount in the range 0.01-1% byweight, preferably in the range 0.03-0.5% by weight and particularlypreferably in the range 0.1-0.3% by weight, based on the weight of thepolymer, is suitable. The catalyst may also be used in an amount of from0.02 to 0.2% by weight.

In one embodiment of the present invention the catalyst can be usedtogether with a co-catalyst. This co-catalyst is preferably a ligand offormula R_(m)B, where R, m and B are as defined above, and m ispreferably 3. Preferably B is phosphorus, and the R groups can be thesame or different. The R group of the catalyst may be a triaryl,trialkyl, tricycloalkyl, diaryl monoalkyl, dialkyl monoaryl, diarylmonocycloalkyl, dialkyl monocycloalkyl, dicycloalkyl monoaryl ordicycloalkyl monoaryl group. Examples of suitable co-catalyst ligandsare given in U.S. Pat. No. 4,631,315, the disclosure of which isincorporated by reference as far as applicable in the respectivejurisdiction. The preferred co-catalyst ligand is triphenylphosphine.The co-catalyst ligand is preferably used in an amount in the range 0 to5000%, more preferably 500 to 3000% by weight, based on the weight ofcatalyst. Preferably also the weight ratio of the co-catalyst to therhodium-containing catalyst compound is in the range 0 to 50, morepreferably in the range 5 to 30.

In one embodiment of the process according to the present invention theunsaturated polymer, hydrogen and the catalyst are introduced into thefirst chamber at the bottom of the multistage agitated reactor

The catalyst may be introduced into one or more different chambers ofthe multistage agitated reactor.

The continuous hydrogenation is preferably carried out at a temperatureof from about 100° C. to about 260° C., preferably from about 100° C. to180° C., most preferably from about 120° C. to about 160° C. and at ahydrogen pressure of from about 0.7 to 50 MPa, most preferably fromabout 3.5 to 10.5 MPa. The continuous hydrogenation is preferablycarried out at a temperature in the range of from 100° C. to 260° C.,preferably in the range of from 100° C. to 180° C., most preferably from120° C. to 160° C. and at a hydrogen pressure in the range of from 0.7to 50 MPa, more preferably in the range of from 3.5 to 10.5 MPa.

The process according to the invention is carried out in a multistageagitated reactor comprising a cylindrical, elongated shell having closedends and separated by baffles into a multiplicity of discrete chamberswith access from one chamber to another through concentric circularopenings, axially centred with said baffles and a continuous rotatableshaft extending concentric with said baffles within said shell with atleast one impeller means attached thereto positioned in each chamber,with said continuous rotatable shaft and said circular openingsproviding annular openings.

Baffles are used to separate the whole reactor into multiple chambers(stages). The baffles have central openings, which allow the access ofboth polymer solution and hydrogen gas. The dimension of the centralopening of the baffle is critical, depending on the reactor scale,polymer solution viscosity and flow rate as well as other reactionconditions, etc., and it can significantly affect the processeffectiveness although it can vary depending on the maximum back mixingallowed within the reactor. For example, for a reactor with a diameterof less than 20″ and with a viscosity of less than 500 cp (correspondingto 0.5 Pas), for a typical polymer hydrogenation system, within ageneral flow rate range the residence time would be 10 mins to a fewhours, the ratio of the diameter of the circular opening to the diameterof the multistage agitated reactor is preferably in the range of from1:10 to 1:2, further preferably from about 1:8 to 1:4.

In a further embodiment the multistage agitaged reactor contains from 3to 30, preferably from 6 to 10, baffles, the baffles being circulardiscs with a central opening, the ratio of the diameter of the centralopening to the diameter of the multistage agitated reactor beingpreferably in the range of from about 1 to 6, further preferable fromabout 1 to 2.

The baffles can preferably be provided, in addition to the centralcircular opening, with some perforations therein.

The number of baffles is typically from 1 to 50, preferably from 3 to30, more preferably from 3 to 20 and most preferably from 6 to 10.

The multistage agitated reactor is typically provided with one or twocontinuous rotatable shafts, extending in the direction of the length ofthe reactor, with at least one impeller attached thereto positioned ineach chamber. The advantage to use two rotatable shafts is that theimpellers on these two shafts could be run at different speeds to meetthe need of the different reactor zones. For example, in the inlet zone,mixing, heat transfer and mass transfer are crucial for which higheragitation speed is preferred, and in the exit zone, extremely highagitation consumes high energy but does not give any significant benefitto compromise the consumed energy. When two rotatable shafts are used,the inlet shaft is preferred to be extended from chamber one to three(“chamber” here means the interval between two baffles), more preferablyfrom chamber one to two, and the other shaft is then extended to theremaining chambers.

The impellers should be able to provide excellent mixing in the lateraldirection and minimal back mixing in axial directions. Depending on theviscosity of the polymer solution, the impellers could be of varioustypes from paddle type, turbine type, propellant type or helicalribbons. For example, for the polymer solution which has a viscosity ofless than 200 cp (0.2 Pas), paddle and turbine types are preferred; thenumber of blades could be preferably four to twelve. When the viscosityis higher than 5000 cp (5 Pas), helical ribbons or a combination ofhelical ribbons and propellant impeller are preferred; the ribbon numbercould be one to four. Furthermore, the impellers along the shaft(s) arenot necessarily the same. In the chambers near the inlet, impellerswhich provide a high shear rate with proper circulation-in-chambercapacity are preferred and in the chambers near the exit, the impellerswhich provide high circulation-in-chamber capacity with proper shearrate are preferred.

The ratio of the diameter of the at least one impeller to the diameterof the multistage agitated reactor is preferably in the range of from19:20 to 1:3, depending on the viscosity of the polymer solution. Forexample, the ratio is preferably from 3:4 to 1:3, further preferablyfrom 1:2 to 1:3. This in particular applies for a system with aviscosity of less than 200 cp (0.2 Pas).

A longer reactor allows a thinner reactor wall under pressure conditionscompared to a single CSTR. Therefore, the cost of reactor material canbe reduced. The continuous hydrogenation reactor also possessesadvantages over a bubble column since an impeller is embedded in eachstage. Therefore, superior mass transfer and heat transfer can beobtained. The superior mass transfer permits a lower hydrogen flow rateand greatly reduces the total amount of excess hydrogen required.

The dimension of the reactor could vary according to the requirement forthe yield and for the residence time. The absolute residence timedepends on the catalyst activity, the catalyst concentration and thehydrogenation degree desired, and the relative residence time ispreferably 3 to 6 times of the kinetic reaction time constant.

FIG. 1 illustrates an example of the multistage agitated reactor with acylindrical, elongated shell (1) having closed ends and separated bybaffles (2) into a multiplicity of discrete chambers with access fromone chamber to another through concentric circular openings, axiallycentered with said baffles and a continuous rotatable shaft (3)extending concentric with said baffles within said shell with at leastone impeller (4) attached thereto positioned in each chamber. In theexemplary multistage agitated reactor (1) the hydrogenated polymer iswithdrawn over a conduit (7) which is located at the opposite end of themultistage agitated reactor to where the feed, i.e. the reactants (5)and hydrogen gas (6) is introduced, wherein the conduit (6) is attachedto a gas sparger (8).

The multistage structure can be advantageous for temperature controlflexibility as the temperature in each chamber can be controlledseparately so that isothermal hydrogenation can be realized.

An advantageous option is that a pre-mixer can be used before thereactor to provide considerable flexibility of the reactor performance.When a pre-mixer is used, the rapid heating and then the instantaneousmixing between the catalyst and polymer solution can be achieved beforethe MAC which can simplify the design of the MAC. The pre-mixer ispreferably a cylinder tank equipped with an agitator having superiormixing performance. The cylinder tank could be disposed vertically orhorizontally. The volume of the pre-mixer is preferably between 1% and100% of the volume of the MAC, depending on the scale of the MAC. Forexample, when the volume in the scale is larger than 100 L the volumeratio of the pre-mixer and the MAC is preferably less than 20%, andfurther preferably less than 10%. The ratio of the length and thediameter is preferably from 0.5 to 3.0, more preferably from 0.5 to 1.0when the pre-mixer is disposed vertically and from 1.0-3.0 when thepre-mixer is horizontally disposed. The pre-mixer can have one ormulti-agitators, depending on the volume of the pre-mixer and also theway of the pre-mixer disposal (vertical or horizontal). Preferably, theagitator(s) in the pre-mixer is a high-shear type of agitator, such as apitched blade agitator or turbines when the pre-mixer is disposedvertically, and turbines or deformed discs when the pre-mixer ishorizontally disposed. A deformed disc here is such a disc impellerwhich is formed by scissor-cutting 12 or 16 lines evenly from the edgetoward the center of the disc and the length of the lines is ⅓ to ⅖ ofthe diameter of the disc and by then twisting each such formed petal bya 30-60 degree in opposite directions for the adjacent petals. The ratioof the agitator diameter and the inner diameter of the pre-mixer ispreferably 1/3-19/20, depending on the viscosity and the disposed mannerof the pre-mixer. Preferably, for example, when the viscosity is lessthan 200 cp (0.2 Pas) and the pre-mixer is vertically disposed, thediameter ratio is from 1/3 to 2/3, and when the viscosity is less than200 cp (0.2 Pas) and the pre-mixer is horizontally disposed, thediameter ratio is larger than 2/3.

The further flexibility of the multistage structure is that a pre-mixerbefore the MAC can be used or can be eliminated. When a pre-mixer isused, the rapid heating and then the instantaneous mixing between thecatalyst and polymer solution can be achieved before the MAC andsimplify the design of the MAC and one shaft would suffice; however, inthe alternative, the first chamber of the MAC can function as thepre-mixer, and a physically separated pre-mixer can be eliminated.Depending on the scale of the process, for a small scale process, thepre-mixer can be eliminated; however, for a large scale process, the useof a pre-mixer is economically beneficial.

In one embodiment of the process according to the invention theunsaturated polymer and hydrogen are passed via a pre-mixer prior tointroducing into the multistage agitated reactor

If a pre-mixer is used, the catalyst can be partly or completelyintroduced via the pre-mixer.

Preferably, heat transfer means are provided to control the temperaturein the reactor and the premixer. Such a type of heat transfer meanscould be a jacket equipped to the MAC and the pre-mixer, in which theheating/cooling medium could be steam or other fluids to designed forheat exchange, or could be a set of coils located inside the reactor orthe pre-mixer, or could be an electrical heating system, and also couldbe the combination of the mentioned means. For a small scale operation(e.g., lab scale), using only one type of heat transfer means ispreferred and internal heating/cooling coils are not needed. However,because the hydrogenation operation is highly exothermic, for a mediumor large scale operation (e.g., pilot to commercial scale), acombination of several heating/cooling means is preferred and internalheating/cooling coils are required.

In one embodiment a cooling coil is provided in the first chamber of themultistage agitated reactor.

Preferably, heat exchange means are provided to cool down the productmixture drawn off from the multistage agitated reactor.

The disposition of the MAC can be vertical, horizontal or in any angle,preferably vertical.

Hydrogen can be introduced into the reaction system from the pre-mixeror the reactor via a gas sparger, in order to ensure uniformdistribution thereof, or from both the pre-mixer and the reactor.

Preferably, the continuous hydrogenation is carried out at a ratio ofthe hydrogen gas flow rate to the flow rate of the unsaturated polymersfrom 0.1 to 100, more preferably from 0.5 to 50, and most preferablyfrom 1 to 10.

Preferably, the liquid residence time in the multistage agitated reactor(1) is from 5 min. to 1 hour, more preferably from 10 min. to 40 min.,and most preferably from 20 to 40 min.

The direction of hydrogen and polymer solution flow through themultistage agitated reactor may be opposite or the same but a paralleloperation manner is preferred. For a vertically or slantwise disposedMAC, the direction of the flow could be upwards or downwards, preferablyupwards. Accordingly, the unsaturated polymer, hydrogen and homogeneouscatalyst are preferably introduced into the first chamber at the bottomof to the multistage agitated reactor, or the catalyst pre-mixer if apre-mixer is used.

The present invention relates especially to the hydrogenation of nitrilerubber.

The term nitrile rubber, also referred to as “NBR” for short, refers torubbers which are copolymers or terpolymers of at least one α,β-unsaturated nitrile, at least one conjugated diene and, if desired,one or more further copolymerizable monomers.

The conjugated diene can be of any nature. Preference is given to using(C₄-C₆) conjugated dienes. Particular preference is given to1,3-butadiene, isoprene, 2,3-dimethylbutadiene, piperylene or mixturesthereof. Very particular preference is given to 1,3-butadiene andisoprene or mixtures thereof. Special preference is given to1,3-butadiene.

As α,β-unsaturated nitrile, it is possible to use any knownα,β-unsaturated nitrile, preferably a (C₃-C₅) α,β-unsaturated nitrilesuch as acrylonitrile, methacrylonitrile, ethacrylonitrile or mixturesthereof. Particular preference is given to acrylonitrile.

A particularly preferred nitrile rubber is thus a copolymer ofacrylonitrile and 1,3-butadiene.

Apart from the conjugated diene and the α,β-unsaturated nitrile, it ispossible to use one or more further copolymerizable monomers known tothose skilled in the art, e.g. α,β-unsaturated monocarboxylic ordicarboxylic acids, their esters or amides. As α,β-unsaturatedmonocarboxylic or dicarboxylic acids, preference is given to fumaricacid, maleic acid, acrylic acid and methacrylic acid. As esters ofα,β-unsaturated carboxylic acids, preference is given to using theiralkyl esters and alkoxyalkyl esters. Particularly preferred alkyl estersof α,β-unsaturated carboxylic acids are methyl acrylate, ethyl acrylate,butyl acrylate, butyl methacrylate, 2-ethylhexyl acrylate, 2-ethylhexylmethacrylate and octyl acrylate. Particularly preferred alkoxyalkylesters of α,β-unsaturated carboxylic acids aremethoxyethyl(meth)acrylate, ethoxyethyl(meth)acrylate andmethoxyethyl(meth)acrylate. It is also possible to use mixtures of alkylesters, e.g. those mentioned above, with alkoxyalkyl esters, e.g. in theform of those mentioned above.

The proportions of conjugated diene and α,β-unsaturated nitrile in theNBR polymers to be used can vary within wide ranges. The proportion ofor of the sum of the conjugated dienes is usually in the range from 40to 90% by weight, preferably in the range from 55 to 75% by weight,based on the total polymer. The proportion of or of the sum of theα,β-unsaturated nitriles is usually from 10 to 60% by weight, preferablyfrom 25 to 45% by weight, based on the total polymer. The proportions ofthe monomers in each case add up to 100% by weight. The additionalmonomers can be present in amounts of from 0 to 40% by weight,preferably from 0.1 to 40% by weight, particularly preferably from 1 to30% by weight, based on the total polymer. In this case, correspondingproportions of the conjugated diene or dienes and/or of theα,β-unsaturated nitrile or nitriles are replaced by the proportions ofthe additional monomers, with the proportions of all monomers in eachcase adding up to 100% by weight.

The preparation of nitrile rubbers by polymerization of theabovementioned monomers is adequately known to those skilled in the artand is comprehensively described in the polymer literature.

Nitrile rubbers which can be used for the purposes of the invention arealso commercially available, e.g. as products from the product range ofthe trade names Perbunan® and Krynac® from Lanxess Deutschland GmbH.

The nitrile rubbers used for the hydrogenation have a Mooney viscosity(ML 1+4 at 100° C.) in the range from 30 to 70, preferably from 30 to50. This corresponds to a weight average molecular weight M_(w) in therange 200 000-500 000, preferably in the range 200 000-400 000. Thenitrile rubbers used also have a polydispersity PDI=M_(w)/M_(n), whereis the weight average molecular weight and M_(n) is the number averagemolecular weight, in the range 2.0-6.0 and preferably in the range2.0-4.0.

Hydrogenated nitrile rubber, also referred to as “HNBR” for short, isproduced by hydrogenation of nitrile rubber. Accordingly, the C═C doublebonds of the copolymerized diene units have been completely or partlyhydrogenated in HNBR. The degree of hydrogenation of the copolymerizeddiene units is usually in the range from 50 to 100%.

Hydrogenated nitrile rubber is a specialty rubber which has very goodheat resistance, excellent resistance to ozone and chemicals and alsoexcellent oil resistance.

The above mentioned physical and chemical properties of HNBR areassociated with very good mechanical properties, in particular, a highabrasion resistance. For this reason, HNBR has found wide use in avariety of applications. HNBR is used, for example, for seals, hoses,belts and clamping elements in the automobile sector, and also forstators, oil well seals and valve seals in the field of oil extractionand also for numerous parts in the aircraft industry, the electronicsindustry, mechanical engineering and shipbuilding.

Commercially available HNBR grades usually have a Mooney viscosity (ML1+4 at 100° C.) in the range from 35 to 105, which corresponds to aweight average molecular weight M_(w) (method of determination: gelpermeation chromatography (GPC) against polystyrene equivalents) in therange from about 100 000 to 500 000. The polydispersity index PDI(PDI=M_(w)/M_(n) where M_(w) is the weight average molecular weight andM_(n) is the number average molecular weight), which gives informationabout the width of the molecular weight distribution, measured here isfrequently in the range from 2,5 to 4,5. The residual double bondcontent is usually in the range from 1 to 18%.

The degree of hydrogenation depends on the polymer concentration, theamount of catalyst used, the gas and liquid flow rates and processconditions. The desired hydrogenation degree is from about 80 to about99.5%, preferably from about 90 to about 99%.

The hydrogenation degree can be determined by using Fourier TransformInfrared (FTIR) or Proton Nuclear Magnetic Resonance (NMR) techniques.

EXAMPLES

The invention is further illustrated by means of examples.

The following examples are set forth to illustrate the scope of theinvention but are not intended to limit the same.

Example 1

A multistage agitated cylinder reactor with an inner diameter of 0.06 mconsists of six stages, vertically disposed. The height of each stage isequal to the diameter of the reactor. Five horizontal baffles separatethe reactor into six stages (chambers) as well as provide a centralopening of 5% of the whole cross sectional area. A Rushton turbine witha diameter of 0.03 m is located in the center of each stage as animpeller. The reactor has a jacket which allows heating/cooling mediumto go through and steam was used as to the heating medium. Thetemperature of the reactor was maintained at 130° C. A pipe nozzle witha diameter of 0.001 m was used as a hydrogen sparger. Nitrile butadienerubber (NBR)(Krynac® 38.5 with 38 wt.-% ACN content and a Mooneyviscosity ML 1+4 at 100° C. of 50) dissolved in monochlorobenzene (MCB)is used as an unsaturated polymer with a 2.5 wt.-% concentration.OsHCl(CO)(O₂)(PC_(y3))₂ was dissolved in monochlorobenzene underhydrogenation conditions as the catalyst precursor. The prepared Osmiumcatalyst concentration was 100 μM. The impeller was operated at 600r.p.m. The system was operated at a pressure of 350 psi (2.41 MPa) by aback pressure regulator. The nitrile butadiene rubber solution wascharged into the reactor at a flow rate of 24 ml/min. Hydrogen gas wassupplied at a flow rate of 48 ml/min.

After the flow became steady, polymer samples were taken at the exit ofthe reactor. The hydrogenation degree of nitrile butadiene rubber wasdetermined by FTIR. The resultant polymer has a hydrogenation level of91%.

Example 2

Example 2 was conducted in the same reactor as described in example 1.The reactor was operated at 140° C. and 500 psig (3.45 MPa) hydrogen.2.5 wt.-% NBR was dissolved in MCB. The liquid feed rate was 24 ml/minand the gas flow rate was 48 ml/min, and the impeller was operated at750 rpm. The same catalyst as in example 1 was used in this example. Theosmium catalyst concentration was 80 μM. The resultant polymer had ahydrogenation level of 99.4% at steady state.

Example 3

Example 3 was conducted in the same reactor as described in example 1.The reactor was operated at 140° C. and 350 psig (2.41 MPa) hydrogen.2.5 wt.-% NBR was dissolved in MCB. The liquid feed rate was 24 ml/minand the gas flow rate was 48 ml/min. The impeller was operated at 750rpm. The same catalyst was used as in example 1. The osmium catalystconcentration was 27 μM. The resultant polymer had a hydrogenation levelof 61% at steady state.

Example 4

Example 4 was conducted in the same reactor as described in example 1.The reactor was operated at 140° C. and 500 psig (3.45 MPa) hydrogen. 5wt.-% NBR was dissolved in MCB. The liquid feed rate was 24 ml/min andthe gas flow rate was 48 ml/min. The impeller was operated at 750 rpm.The same catalyst was used in as in example 1. The osmium catalystconcentration was 140 μM. The resultant polymer had a hydrogenationlevel of 99% at steady state.

Example 5

Example 5 was conducted in the same reactor as described in example 1.The reactor was operated at 140° C. and 500 psig (3.45 MPa) hydrogen.2.5 wt.-% NBR was dissolved in MCB. The liquid feed rate was 48 ml/minand the gas flow rate was 48 ml/min. The impeller was operated at 750rpm. The same catalyst was used as in example 1. The osmium catalystconcentration was 80 μM. The resultant polymer had a hydrogenation levelof 85% at steady state.

The above 5 examples demonstrate that the present invention effectivelyprovides hydrogenation of unsaturated polymers and exhibits operationflexibility.

1. A process for the continuous hydrogenation of carbon carbon doublebonds in an unsaturated polymer based on a conjugated diolefin and atleast one other copolymerizable monomer to produce a hydrogenatedpolymer, in the presence of a solvent and a catalyst, wherein saidunsaturated polymer, hydrogen and said catalyst are passed through amultistage agitated reactor, wherein said multistage agitated reactorcomprises a cylindrical, elongated shell having closed ends andseparated by baffles into a multiplicity of discrete chambers withaccess from one chamber to another through concentric circular openings,axially centered with said baffles and one or two continuous rotatableshafts extending concentric with said baffles within said shell with atleast one impeller attached thereto positioned in each chamber, saidhydrogenated polymer being withdrawn at the opposite end of themultistage agitated reactor at which the feed is introduced.
 2. Theprocess according to claim 1, wherein said catalyst is either ahomogeneous catalyst or a heterogeneous catalyst.
 3. The processaccording to claim 2, wherein said homogeneous catalyst is anorgano-metallic catalyst.
 4. The process according to claim 3, whereinsaid homogeneous catalyst is a rhodium, ruthenium, osmium, or iridiummetal complex catalyst.
 5. The process according to claim 2, whereinsaid heterogeneous catalyst comprises one or more of the metalsplatinum, palladium, nickel, copper, rhodium and ruthenium.
 6. Theprocess according to claim 5, wherein said heterogeneous catalyst issupported on carbon, silica, calcium carbonate or barium sulphate. 7.The process according to claim 3, wherein the organo-metallic catalystis a rhodium or ruthenium metal complex catalyst having the formula(R_(m) ¹B)_(l)M X_(n), where M is ruthenium or rhodium, the groups R¹are identical or different and are each a C₁-C₈-alkyl group, aC₄-C₈-cycloalkyl group, a C₆-C₁₅-aryl group or a C₇-C₁₅-aralkyl group, Bis phosphorus, arsenic, sulphur or a sulphoxide group S═O, X is hydrogenor an anion, l is 2, 3 or 4, m is 2 or 3 and n is 1, 2 or
 3. 8. Theprocess according to claim 7, wherein X is halogen.
 9. The processaccording to claim 8, wherein X is chlorine or bromine.
 10. The processaccording to claim 7, wherein the catalyst is selected from the groupconsisting of tris(triphenyl-phosphine)rhodium(I) chloride,tris(triphenylphosphine)rhodium(III) trichloride,tris(dimethylsulphoxide)rhodium(III) trichloride,tetrakis(triphenylphosphine)-rhodium hydride of the formula(C₆H₅)₃P)₄RhH and the corresponding compounds in which thetriphenylphosphine has been completely or partly replaced bytricyclohexylphosphine.
 11. The process according to claim 3, whereinthe organo-metallic catalyst is an osmium metal complex catalyst havingthe formulaOs QX(CO) (L) (PR₃)₂ in which Q may be one of hydrogen and a phenylvinylgroup, X may be one of halogen, tetrahydroborate and alkyl- oraryl-carboxylate, L may be one of an oxygen molecule, benzonitrile or noligand, and R may be one of cyclohexyl, isopropyl, secondary butyl andtertiary butyl said tertiary butyl being present only when one R ismethyl, with the proviso that when Q is phenylvinyl X is halogen and Lis no ligand and when X is alkyl- or aryl-carboxylate Q is hydrogen andL is no ligand, said halogen being selected from chlorine and bromine.12. The process according to claim 1, wherein a co-catalyst is alsopresent.
 13. The process according to claim 12, whereintriphenylphosphine is used as co-catalyst.
 14. The process according toclaim 1, wherein said unsaturated polymer, hydrogen and said catalystare introduced into the first chamber at the bottom of the multistageagitated reactor.
 15. The process according to claim 1, wherein thecatalyst is introduced into one or more different chambers of themultistage agitated reactor.
 16. The process according to claim 1,wherein said conjugated diolefin is one or more substances selected frombutadiene, isoprene, piperylene and 2,3-dimethylbutadiene.
 17. Theprocess according to claim 1, wherein at least one other copolymerizablemonomer is one or more substances selected from acrylonitrile, propylacrylate, butyl acrylate, propyl methacrylate, methacrylonitrile, butylmethacrylate and styrene.
 18. The process according to claim 1, whereinthe catalyst is introduced in a total quantity of from 0.01 to 1.0 wt.-%based on the unsaturated polymer.
 19. The process according to claim 18,wherein the catalyst is introduced in a total quantity of from 0.03 to0.5 wt.-%, based on the unsaturated polymer.
 20. The process accordingto claim 18, wherein the catalyst is introduced in a total quantity offrom 0.1 to 0.3 wt.-% based on the unsaturated polymer.
 21. The processaccording to claim 1, wherein the continuous hydrogenation is carriedout in the presence of a hydrocarbon solvent.
 22. The process accordingto claim 21, wherein the continuous hydrogenation is carried out in thepresence of a hydrocarbon solvent selected from the group consisting ofbenzene, toluene, xylene, monochlorobenzene and tetrahydrofuran.
 23. Theprocess according to claim 22, wherein the concentration of theunsaturated polymer in the hydrocarbon solvent is from about 1 to about40 wt.-%.
 24. The process according to claim 23, wherein theconcentration of the polymer in the hydrocarbon solvent is from about 2to about 20 wt.-%.
 25. The process according to claim 1, wherein thecontinuous hydrogenation is carried out at a ratio of the hydrogen gasflow rate to the flow rate of the unsaturated polymers of from 0.1 to100.
 26. The process according to claim 25, wherein the continuoushydrogenation is carried out at a ratio of the hydrogen gas flow rate tothe flow rate of the unsaturated polymers of from 0.5 to
 50. 27. Theprocess according to claim 25, wherein the continuous hydrogenation iscarried out at a ratio of the hydrogen gas flow rate to the flow rate ofthe unsaturated polymers of from 1 to
 10. 28. The process according toclaim 1, wherein the residence time of the solvent in the multistageagitated reactor is from 5 min. to 1 hour.
 29. The process according toclaim 28, wherein the residence time of the solvent in the multistageagitated reactor is from 10 min, to 40 min.
 30. The process according toclaim 1, wherein the continuous hydrogenation is undertaken at atemperature of from about 100° C. to about 260° C.
 31. The processaccording to claim 1, wherein the continuous hydrogenation is carriedout at a hydrogen pressure of from about 0.7 to 50 MPa.
 32. The processaccording to claim 1, wherein said unsaturated polymer and hydrogen arepassed via a pre-mixer prior to introducing into the multistage agitatedreactor.
 33. The process according to claim 1, wherein the catalyst ispartly or completely introduced via a pre-mixer.
 34. The processaccording to claim 1, wherein hydrogen is introduced via a gas sparger.35. The process according to claim 1, wherein a cooling coil is providedin the first chamber of the multistage agitated reactor.
 36. The processaccording to claim 1, wherein heat exchange meanings are provided tocool down the product mixture drawn off from the multistage agitatedreactor.
 37. The process according to claim 1, wherein there areprovided from 3 to 30 baffles, the baffles being circular discs with acentral opening, the ratio of the diameter of the central opening to thediameter of the multistage agitated reactor being in the range of fromabout 1 to
 6. 38. The process according to claim 37, wherein there areprovided from 6 to 10 baffles and the ratio of the diameter of thecentral opening to the diameter of the multistage agitated reactor beingin the range of from about 1 to
 2. 39. The process according to claim 1,wherein the ratio of the diameter of the at least one impeller to thediameter of the multistage agitated reactor is in the range of from 3:4to 1:3.